Processes for carrying out catalytic endothermic high-pressure gas reactions

ABSTRACT

Processes for carrying out catalytic exothermic and endothermic high-pressure gas reactions with a single-walled pressure vessel or shell containing cross-flow (e.g., radial flow) heat transfer exchangers, a continuous catalytic bed having at least two stages, and means for effecting &#34;cross-over&#34; material flows from &#34;outside&#34; to &#34;inside&#34; (for exothermic reactions) and vice versa (for endothermic reactions), whereby conditions of: maximum gas temperature always being in the core of said vessel or shell, minimal pressure drop, and minimal compression of catalyst particles are achieved, along with significant economic savings in cost of the pressure vessel or shell and catalyst (through extension of catalyst life).

This is a continuation of application Ser. No. 285,229, filed on July20, 1981, abandoned, which is a division of Ser. No. 130,895, filed onMar. 17, 1980, abandoned, which is a division of Ser. No. 41,378, filedMay 22, 1979, now U.S. Pat. No. 4,341,737.

BACKGROUND OF THE INVENTION

1. Field of the Invention

This invention relates to an improved high-pressure process foreffecting exothermic or endothermic gaseous reactions such that maximumgas temperatures are always in the core of the apparatus utilized forsuch process and minimum pressure drop conditions and significanteconomies can be achieved, resulting in extension of catalyst life and amarked decrease in the capital cost of such apparatus.

More specifically, this invention relates to an improved high-pressureprocess, capable of effecting either exothermic or endothermic gaseousreactions with the foregoing results which utilizes a single-walledpressure vessel or shell containing cross-flow, as, e.g., radial flow,heat transfer exchangers, a continuous particulate catalyst bed havingat least two stages, and means for effecting "cross-over" material flowssuch that, for exothermic reactions, material flows are radiallydirected inwardly from "outside" to "inside", whereas, for endothermicreactions, material flows are radially directed outwardly from "inside"to "outside".

2. Description of the Prior Art

Heretofore, the art has been replete with high-pressure processes andcatalytic reactors or converters for effecting the gaseous syntheses ofsuch valuable materials as ammonia, methanol, hydrocyanic acid,hydrogen, methane, and styrene. Typically, such apparatus have had to bebuilt to withstand the extreme pressures and temperatures associatedwith such syntheses, approximating wide limits varying, for example,from between 1200 to 10,000 p.s.i.g. Thus, in order to accommodate thecommercial production rates required, e.g., catalytic converters capableof generating 1000 tons of ammonia per day, double-walled reactorvessels of enormous size have had to be employed as shown in U.S. Pat.No. 3,567,404. However, the costs and difficulties of manufacturing suchconverters have likewise been enormous. Moreover, equipment sizingproblems have also been encountered, since in order to maintain spaceand linear velocity conditions at reasonable pressure drops, convertersof prohibitively large diameters, in view of their high operatingpressures, are required. Furthermore, it is well-known in the art that,for a given operating pressure and temperature, the larger the diameterof the vessel, the thicker its walls have to be. Since the materials ofvessel construction are also influenced by the temperature as well as bythe hydrogen partial pressure, the reason for use of conventionaldouble-walled vessels in the part has been manifest.

Accordingly, the art has long been concerned with providing reactors orconverters of increased production for high-pressure processes suitablefor large-scale reactions, within the limits of acceptable designcriteria and having flow patterns of reactants which lend themselves toincreased production through increased length of the reactor orconverter rather than through an increase in such reactor's orconverter's diameter.

This too has posed problems in view of the fact that, in order toaccommodate the increased production requirements for such processes,tall reactors or converters on the order of 40-50 feet high arerequired. Since within such reactors or converters one or more beds ofcatalytic contact material has to be vertically disposed, maintenance ofoptimum space and linear velocity conditions without prohibitivepressure drops has not been attainable, and various means have beensought to solve this problem.

One such solution, with respect to such process deficiencies, has beenproposed in U.S. Pat. No. 3,567,404, which utilizes the conventionaldouble-walled reactor, whereby the reactant gases are permitted to flowin a directional perpendicular to the longitudinal axis of the outershell and the inner reaction zone, and across one or more catalyst bedsin series, such that the gases flow from one bed to the next consecutivebed through a passageway therebetween, the direction of flow of thegases through said passageway being generally opposite to theirdirection of flow through the catalyst bed. The arrangement of flow inthis manner greatly facilitates the manner in which the reaction isconducted and permits wide alteration of desirable variables. Forexample, by having reactant flow downward across one bed and upwardthrough an adjacent bed, this flow pattern has the effect of shorteningthe converter by eliminating the passageways between the beds. However,such flow methods and patterns have been unsuccessful because they havebeen unable to satisfy the temperature requirements associated withoptimal yields and maximum suppression of competing side reactions,notwithstanding the use of heat exchange means disposed to accommodatesuch flow methods and patterns. Moreover, these flow patterns aresubject to increased flow resistance, thereby leading to increasedpressure drops, and a considerably reduced circulation rate through thereactor for a given catalyst volume. The solution to this type offlow-type, process problem has been the adoption of radial flow meanssuch as taught by (1) U.S. Pat. No. 3,372,988, which originated the ideaof "means for passing a synthesis gas through the catalyst bodiessuccessively in opposite radial directions"; and by (2) an improvedversion of radial flow in U.S. Pat. No. 3,472,631 whereby the reactantgases are made to flow through each successive catalyst bed layer moreor less horizontally in the reverse direction to the preceding catalystlayer and around heat exchange tubes at turning points counter-currentto the fresh reactant gases.

Additionally, the concept of circulating feed gas through tubes disposedin the catalyst bed for cooling purposes, prior to actual contact of thefeed gas with the catalyst, has been shown in U.S. Pat. Nos. 2,853,371;3,041,161; 3,050,377; and 3,212,862. The alternative approach to thismode of cooling has been through the use of quench cooling andquench-type converters as shown in U.S. Pat. Nos. 2,495,262; 2,632,692;2,646,391; 3,366,461; 3,396,685; 3,433,600; 3,433,910; 3,458,289;3,475,136; 3,475,137; 3,498,752; and 3,663,179. In the prior artquench-type converters, the quench fluid has generally been added to themain reactant stream between separate beds consisting of solid catalystgranules, spheres, or the like. The quench-cooled apparatus, however,have suffered from the disadvantages of high pressure drop and increasedcost and complexity.

Heretofore, however, none of the prior art high-pressure processes, orcatalytic apparatus utilized therein, for performing reactions in thegaseous phase have been effective for both exothermic and endothermicreactions; none have known that it was possible to use single-walledapparatus adapted to accommodate radial material flow patternsintegrated with cross-flow heat exchangers. Moreover, none of the knownapparatus have been successful at maintaining maximum gas temperaturesin the core of the apparatus with minimum pressure drop conditionswithin the limits of acceptable design criteria and acceptable flowpatterns of reactions. In particular, and most notable is the fact thatthe prior art has been unaware of the use of cross-flow, e.g., radial,heat exchangers (let alone of their use with single shell reactors),thereby to enable flow direction to be arranged to make gas expansion orcontraction consistent with catalyst cross-section expansion orcontraction. The present invention has been developed to fill this void,and it does so through means of processes which employ a newconceptually-based design of apparatus which enables the use of asingle-walled reactor or reactor system having multiple reaction stages,whereby a radial flow of reactants is developed in accordance with a"cross-over" pattern such that material flows are directed from"outside" to "inside" for exothermic reactions and vice versa forendothermic reactions. For exothermic reactions, one form of the presentprocess provides one-cross-flow heat transfer stage for each reactionstage. However, in another form of the present process concerned withendothermic reactions, the first heat transfer stage is external to thesystem (for example, it can be situated outside the reactor), and hencefor such endothermic reactions, there are one fewer heat transfer stagesthan reactor stages.

SUMMARY OF THE INVENTION

In accordance with the present invention, there is provided ahigh-pressure process for carrying out either catalytic exothermic orendothermic reaction in the gaseous phase, wherein a single orsingle-walled pressure containment vessel or shell is utilized forconducting such reactions therein in multiple stages, such pressurevessel or shell containing one or more cross-flow (e.g., radial flow)heat transfer exchangers and a continuous catalyst bed having aplurality of stages (at least two), with a cross-flow heat transferstage after each reaction stage for exothermic reactions and one lesscross-flow heat transfer stage than there are reaction stages forendothermic reactions, the arrangement of these stages being such as toprovide "cross-over" material flows in a continuous, uni-directionalflow path through the various catalysts and heat exchanger stages, suchuni-directional flow path being from "outside" to "inside" forexothermic reactions and vice versa for endothermic reactions, therebyeffectively maintaining the highest gaseous temperature in the core ofthe apparatus rather than on the exterior or pressure containment wallsthereof and also minimizing pressure drop.

DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a block-flow schematic diagram, illustrative of the inflowpattern of a typical two-stage, exothermic reaction process such aswould be useful in the synthesis of ammonia, methanol, and methane.

FIG. 2 shows a cross-sectional elevation of the high-pressure apparatusutilized in the pressure processes, wherein the continuous particulatecatalyst bed is depicted in the form of stacked stages, and the variousinflow feed and unreacted reactant/reaction effluent, flows are alsoshown.

FIG. 3 depicts an alternative embodiment of the apparatus depicted inFIG. 2 and shows a cross-sectional elevation of the presenthigh-pressure apparatus wherein the continuous particulate catalyst bedis depicted in the form of concentric stages, and the various inflowfeed and unreacted reactant/reaction effluent flows are also shown.

FIGS. 4 and 5 depict the outflow flow patterns of a typical three-stageendothermic outflow process, with FIG. 4 showing a block-flow diagramthereof and FIG. 5 showing a cross-sectional elevation of thehigh-pressure apparatus used in one embodiment of the present processwherein the particulate catalyst bed (represented in the form of stackedstages) and the various outflow unreacted reactant/reaction effluentoutflows are also shown. For both endothermic systems, there is one lessheat transfer stage than there are reaction stages.

FIGS. 6 and 7 depict the outflow flow patterns of a typical four-stageexothermic flow process, with FIG. 6 showing a block-flow diagramthereof and FIG. 7 showing a cross-sectional elevation of the presenthigh-pressure apparatus used in one embodiment of the present process,wherein the particulate catalyst bed is depicted in the form of eitherstacked or concentric stages, and the various outflow unreactedreactant/reaction effluent outflows are also shown.

FIG. 8 depicts a simplified cross-sectional elevation of a two-stageconverter useful, in accordance with the process of the presentinvention, for ammonia synthesis.

FIG. 9 is a cross-sectional plan view of the ammonia converter shown inFIG. 8 taken along the line A--A thereof.

FIG. 10 is a cross-sectional plan view of the inter-stage cross-overstructure of the ammonia converter shown in FIG. 8 taken along the linesB--B thereof.

And finally, FIG. 11 is a schematic flow-diagram, the temperatureprofile of which is discussed below, of a typical ammonia synthesisprocess conducted in the present catalytic reactor of converter.

DESCRIPTION OF THE DRAWINGS AND PREFERRED EMBODIMENTS

Referring now to FIGS. 1-3, a cold, fluid feed stream 1, which, in apreferred embodiment of the present process, would typically consist ofan ammonia or methanol synthesis gas, is divided into a plurality ofstreams, of which two (2,3) are main feed streams and two (4,5) areby-pass streams. Main feed streams 2 and 3 are first passed,respectively, through the tube side of heat exchangers 6 and 7, wherethey become, respectively, heated streams 8 and 9 which are then joinedinto one main heated stream 10 which is passed into the first reactionstage 11 of a vertically oriented reactor or converter vessel andreacted under the required reaction conditions, which are well-known inthe art. From reactor first stage 11, an effluent stream 12 is passedthrough the shell side of heat exchanger 7 where its heat of reaction isexchanged with the cold main feed stream 3.

The cooled effluent 13 is introduced into the second stage 14 of thereactor, further reacted, and the resultant effluent therefrom 15 ispassed through the shell side of the heat exchanger 6, where the heat ofreaction effluent 15 is exchanged with the cold main feed stream 2. Theresultant cooled effluent stream 16 is discharged from the reactor fordownstream processing.

It is to be noted, with reference to FIGS. 1-3, that (1) flow directionis radially inward, thereby forcing the coolest gases within thisexothermic system outward and that (2) temperature control is achievedby passing the cold fluid feed streams 1, 2, or 3 into streams 10 and13.

The uni-directional flow patterns schematically depicted in FIG. 1 canperhaps better be visualized by recourse to FIGS. 2 and 3, which tracethe course of a cold feed fluid stream to and from the two reactorstages and via the heat exchangers.

Similarly, by recourse to FIGS. 4-7, the uni-directional radial flowpatterns of the fluid feed stream inflows and effluent inflows andoutflows through the various reactor and/or heat exchanger stages can beseen with reference to other embodiments within the purview of theprocesses of the present invention such as, e.g., the use of concentricstages, a three-stage endothermic system (with particular emphasis onoutflows); a four-stage exothermic system, etc. wherein the first twostages are as previously described with reference to FIGS. 1-3 and thethird and fourth stages of the reactor are denoted 17 and 25; thevarious reactor effluent streams are denoted 18, 20, 22, and 26; and thevarious heat exchanged (cooled) reactor effluent streams are denoted 19,21, 24, and 28; and the third-stage, and fourth-stage, heat exchangersare denoted, respectively, as 23 and 27.

Referring now to FIG. 8, there is shown a preferred embodiment of theprocesses of the present invention, which depicts a two-stage ammoniasynthesis converter for use therein. The process flows are as previouslydescribed with reference to FIGS. 1-7, particularly FIGS. 1-3.

A cold fluid stream 1, typically ammonia synthesis gas, is divided intoa plurality of streams comprising two main feed streams 2 and 3, each ofwhich is passed through a low pressure differential diffuser and thenthrough the tube side of cross-flow heat exchangers 6 and 7, where theybecome, respectively, heated streams 8 and 9 which are then joined intoone main heated stream 10 which is passed into the first stage 11(comprising a bed of a conventional, active, particulate catalyst) of avertically oriented reactor or converter vessel and reacted under therequired, conventional reactio conditions. From the reactor first stage11, an effluent stream 12 is passed through the shell side of heatexchanger 7 where its heat of reaction is exchanged with the cold mainfeed stream 3.

The cooled effluent 13 is introduced into the second stage 14 of thereactor, further reacted, and the resultant effluent therefrom 15 ispassed through the shell side of the heat exchanger 6, where the heat ofreaction of effluent 15 is exchanged with the cold main feed stream 2.The resultant cool effluent stream 16 is discharged from the reactor fordownstream processing.

Preferably, the reaction stages include perforated plates and screens 30to facilitate catalyst retention as well as passage of gaseous reactantand product flows into and through the catalyst beds. Furthermore, it isalso preferred that the annular catalyst containers 31 also beperforated so as to minimize pressure differentials and eliminate flowgradiants.

Referring to FIG. 9, this shows a plan view of the interior of thereactor, specifically depicting the cross-flow exchanger 7, the activecatalyst bed of the first reactor stage 11, the product effluent 15 ofthe second reaction stage, the final cooled product stream 16, and thecatalyst retaining cylinder.

Referring to FIG. 10, this shows the heated reactant feed streams 8 and9 for the first reaction stage and the exit patterns for inter-stagecross-over of the cooled product effluent 13 from said first reactionstage.

And, finally, referring to FIG. 11, this depicts a schematic flowdiagram for a conventional two-stage, exothermic reaction process forammonia synthesis, for which diagram a typical temperature profile ofthe various reactant or product flows is discussed below. For example,as shown in FIG. 11, the cold, fluid feed stream 1 is divided into twomain streams 2 and 3. Feed streams 1, 2, and 3 are at a temperature ofabout 350° F. The flow of the other part 3 of feed stream 1, controlledby a valve, after passage through heat exchanger 7, attains atemperature of about 750° F., and is denoted as stream 9. Streams 8 and9, both at about 750° F., unite into combined stream 10 which is at atemperature of about 750° F. and is fed into an active catalyst bed 11.The effluent 12 from the bed is at a temperature of about 954° F.Effluent 12, after being passed through heat exchanger 7, emerges asstream 13 at a temperature of about 735° F. and proceeds into the nextreaction stage, active catalyst bed 14, from which the effluent 15emerges at a temperature of about 885° F.

A temperature profile, such as is discussed above for FIG. 11, thatwould be typical for a process using a 1500 MTD ammonia converter, suchas shown in FIG. 8, is given in the Table 1 below, along with streamflow rates (defined according to the standard ACFS classification). Thestreams are numerically set forth in Table 1, as represented in FIG. 8.

    ______________________________________                                        Stream Number                                                                             Temperature, °F.                                                                    Flow Rate (ACFS)                                     ______________________________________                                        1           350          46.5                                                 2           350          20.8                                                 3           350          25.7                                                 8           750          31.2                                                 9           750          38.6                                                 10          750          69.8                                                 12          954          75.5                                                 13          735          63.8                                                 15          885          67.1                                                 16          703          58.4                                                 ______________________________________                                    

In many high-pressure processes using catalytic converters or reactors,and in all ammonia and methanol processes using such apparatus, there isa pressure containment problem which, under conventional conditions,becomes aggravated by changes in flow direction which increase theamount of pressure drop. A large part of the pressure containmentproblem arises from the fact that the usual feedstocks enter the reactorrelatively cool and exit from it in a relatively hot state. For example,in the ordinary ammonia synthesis process, the synthesis feed isintroduced into the reactor at relatively cool temperaturesapproximating 750° F., and the product effluent exits therefrom at arelatively hot temperature of 950° F. With the present processes, theouter or pressure containment walls of the apparatus utilized need onlybe the thickness of a single wall because, through the use ofuni-directional material flow patterns and the use of a system ofcross-flow (e.g., radial flow) heat exchangers integrated with thecorresponding system of reaction stages to promote such flow patterns,the coolest and lowest possible temperatures in the reactor are at theouter or pressure containment wall.

In common industrial practice, the number of reaction stages reaches aneconomic optimum very rapidly, e.g., in about two or three stages, owingto limitations in catalyst activity, the build-up of pressure drop, andthe consequent increase in horsepower for the bulk transport of thegases flowing through such a system. And further complications areintroduced in achieving optimal flow and temperature control throughoutthe system as the number of stages increases.

Application of the present processes to these practical problems enablesthe operator to realize increased yields, increased conversions,increased thermal efficiency, and decreased pressure drop.

In the most preferred or best mode embodiment of the processes of thepresent invention, "cross-flow" tubular heat exchangers are integratedwith catalyst-containing annuli such that the shell-side fluid in theheat exchanger flows radially in a direction substantially normal tothat in which the reactant fluid flows. However, the geometry of thearrangement between the radial-flow heat exchangers and thecatalyst-containing annuli is such that the respective cooling fluid ofthe heat exchangers and the fluid comprising the mixture of unreactedfeedstock and product effluent in the catalyst tubes both follow acontinuous uni-directional path in inter-related patterns, whereby theproduct effluent of each reaction stage is cooled in a subsequent heatexchange stage.

It is especially preferred that, for each exothermic reaction stage, oneannular particulate catalyst bed and one cross-flow heat exchanger isprovided (one less cross-flow heat transfer stage, as previously noted,being required for endothermic reactions), and this arrangement isamenable to a variety of operable forms.

For example, the heat exchangers and reaction stages could be positionedside-by-side in stacked vertical formation, with the reaction stageslocated closer to the pressure containment wall. In such an embodiment,as shown in FIG. 2, relative to an exothermic process, a cool mainstreamwould be fed to the lower heat exchanger and heated to reactiontemperature, then passed from "outside" in a plurality of radiallyflowing streams to "inside" through both the upper reaction and exchangestages and then from "outside" to "inside" through the lower reactionand exchange stages.

In most preferred embodiments of the processes of this invention, thereactants are dispersed radially through the various catalyst beds andradially through the heat exchangers, and the cooling fluid of the heatexchangers passes through the exchanger tubes essentially normal to theflow of the reactants.

The invention will now be further illustrated by reference to thefollowing specific, but non-limiting, examples.

EXAMPLE 1

This example illustrates the influence on catalyst life that reactantmaterial fluid flows have in high pressure processes, and compares thecontainer volume changes for an inflow process design and outflowprocess design under the conditions proposed for practice in accordancewith the present invention relative to a process for the synthesis ofammonia by the Haber process, operated at about 950° F. with a cold feedat about 750° F.; at reaction pressures of about 3000 psi; and a partialpressure of hydrogen of about 2300 psi; and a typical volumetricexpansion (molal and thermal) of about 5-8 percent.

As shown in FIGS. 2, 8, and 9, relative to Example 1, the inner diameterof the heat exchange shell is D₁ and approximates 35"; the innerdiameter of the catalyst annulus is D₂ and approximates 80"; and theinner diameter of the reactor itself is D₃ and approximates 90".

In an annular bed of catalyst under conditions of radial flow (withradial flow heat exchanger), the catalyst container volume expands morerapidly than the bulk catalyst as the temperature of the system israised. For a given thermal cycle, i.e., the time from which thecatalyst is loaded into the catalyst container at atmospherictemperature, the reactor is operated at reaction conditions (whereby thecontainer volume increases relative to that of the catalyst and thecatalyst physically settles), and the system cooled to ambienttemperature (whereby the container contracts and compresses the bulkcatalyst, thereby crushing some catalyst particles and causing somedegree of catalyst attrition), the computation of the container volumechange--assuming an ambient temperature of 80° F. and D₁ being stainlesssteel Type 304 for inflow design and ferritic steel (21/4 Cr-1 Mo) foroutflow design; D₂ being ferritic steel (21/4 Cr-1 Mo) for outflowdesign and stainless steel Type 304 for inflow design; and therespective coefficients of the thermal expansion for Type 304 being10.2×10⁻⁶ in/in/°F. and for 21/4 Cr-1 Mo being 7.5×10⁻⁶ in/in/°F.--canbe calculated as follows:

    __________________________________________________________________________     ##STR1##                                                                     Inflow Process Design  Outflow Process Design                                 __________________________________________________________________________    D.sub.1 = (950 - 80) × 10.2 × 10.sup.-6 × 35                                       (750 - 80) × 7.5 × 10.sup.-6 ×                              35 = 0.176"                                            D.sub.2 = (750 - 80) × 7.5 × 10.sup.-6 × 80                                        (950 - 80) × 10.2 × 10.sup.-6 ×                             80 = 0.710"                                             ##STR2##                                                                                             ##STR3##                                                ΔV = 28.458 - 28.225 = 0.233 ft.sup.3 % =                                                    28.780 - 28.225 = 0.555 ft.sup.3 = 1.966%              0.233/28.225 = 0.825%                                                         __________________________________________________________________________

For purposes of this example, the conventional ammonia catalyticconverter design for the process of this invention would be taken as amodel in which the catalyst would be placed in a separate container,usually denoted as a "basket", which would be mounted concentricallywithin a standard pressure vessel used for ammonia synthesis, and inwhich model the annulus between the two vessels would contain the coldfeed gas, thereby providing the option for designing the pressurecontainer to be useful in a non-critical temperature range.

For axial flow reactors, this option would not be available since thecatalyst container would be exposed to the maximum gas temperature.However, this option would be available for radial flow reactors such asthose of the present invention wherein the operator would have theoption of having 750° F. at the catalyst container wall via an "inflow"design or 950° F. through the use of an "outflow" design.

Under the conditions of Example 1, the following calculationdemonstrates there to be about a 25 percent increase in containment costfor the outflow process relative to the inflow process even withoutconsideration of the increased complexity attributable to the presenceof the basket. This calculation is based on the conventionalrequirements of an ASME Section VIII, Division 2 design for a pressurecontainer.

    ______________________________________                                        Wall                      Thickness                                           ______________________________________                                        Inflow Design (Gas Flow = 750° F.)                                                             D.sub.2 = 80"                                                                 D.sub.3 = 90"                                          Pressure Containment Wall (P = 3000 psi,                                                               6.75"                                               T = 750° F.)                                                           (Ferritic Steel - 21/4 Cr:1 Mo-A387-CR22C62)                                                            0.375" (min.)                                       Wall B                                                                        Outflow Process Design (Gas Flow T = 950°.)                                                      D.sub.1 = 80"                                                                 D.sub.2 = 90"                                                                 D.sub.3 = 100"                                      Pressure Containment Wall (P = 3030 psi,                                                                7.25"                                               T = 350° F.)                                                           (carbon steel - A516-70)                                                      Basket Wall (External pressure 30 psi,                                                                  1.00"                                               T = 950° F.)                                                           (stainless steel-Type 304)                                                    Wall E                    0.375" (min.)                                       ______________________________________                                    

Assuming present relative costs per pound as the following:

    ______________________________________                                        Carbon Steel = 1.0                                                            21/4 Cr-1 Mo = 1.5                                                            Type 304 = 4.0                                                                The cost of A = 7.5 × 6.75 × 1.5 =                                                         238.6                                                The cost of B = 6.7 × 0.375 × 1.5 =                                                         11.8                                                                         250.4                                                The cost of C = 8.3 × 7.25 × 1.0 =                                                         189.0                                                The cost of D = 7.5 × 1.00 × 4.0 =                                                          94.2                                                The cost of E = 6.7 × 0.375 × 4.0 =                                                         31.6                                                                         314.8                                                 ##STR4##                                                                     ______________________________________                                    

From Example 2, it can readily be seen that the outflow process designresults in a 25 percent increase in containment costs for the outflowdesign over the inflow process design, apart from consideration of theproblems of complexity attributable to the presence of the catalystbasket.

Within the context of costs, a double wall reactor vessel, i.e., apressure vessel containing a reactor basket, has several otherconstraints which increase cost and mechanical difficulty. For example,it is often considered prudent to incorporate a full diameter closurefor the pressure vessel in order to facilitate assembly of the basket orto provide means of achieving direct access for periodic inspections ofthe pressure vessel. Such a closure is quite costly and increasinglydifficult to achieve as diameter increases.

Moreover, because of the significant differential thermal expansionbetween the pressure vessel and the basket, it is not practical toachieve side entry into the reactor; therefore all connections forinstruments and by-passes must enter through the top of the basket, andbe piped into the required locations in the reactor. It is thereforemanifest that neither of the above constraints exists for asingle-walled vessel; for the access opening need only be large enoughto accommodate passage of the heat exchanger tubular bundles;accordingly, side entry into the single-walled reactor is simple anddirect.

As and wherever defined herein, with respect to the present processes,the terminology "single" or "single walled" pressure vessel or shell, or"single" or "single-walled" apparatus, reactor, or converter, or thethickness thereof is meant to denote the conventional meaning suchterminology has in the art as, for example, defined in Section VIIIRules for Construction of Pressure Vessels, Division 2--AlternativeRules; relative to the ASME Boiler and Pressure Vessel Code, An AmericanNational Standard (ANSI/ASME BPV-VIII-2), 1977 Edition, July 1, 1977, ofthe American Society of Mechanical Engineers.

In like manner, the metals of construction of the apparatus utilized inthe processes of the present invention in respect of the exothermic orendothermic reaction conditions under which such apparatus is intendedto operate, as defined herein, are conventionally determinable, as,e.g., from application of the so-called "Nelson Chart" of G. A. Nelsoncontained in "Steels for Hydrogen Service at Elevated Temperatures andPressures in Petroleum Refineries and Petrochemical Plants" of theAmerican Petroleum Institute, API Publication 941, Second Edition, June1977.

As will be apparent, the essential basis of the processes of the presentinvention is predicated, as has been stated or suggested previously,upon achieving, to the greatest extent possible, conditions of minimalpressure drop and compression of catalyst particles, thereby leading toincreased conversions, yields, catalyst life, and thermal efficiency, bymeans of the utilization and deployment of uni-directional process oreffluent flow patterns and of a system of cross-flow, e.g., radial flow,heat exchangers integrated with the corresponding system of reactionstages to promote such flow patterns such that, for exothermicreactions, the uni-directional process or effluent flow proceedsinwardly from the pressure containment walls of the reactor and viceversa for endothermic reactions. As will also be apparent, the processesof the present invention provide the potential for achieving largeproduction capacity within a single reactor without increasingengineering complexity and without requiring new art or techniques forwelding or pressure vessel fabrication. Accordingly, such importantparameters as nature and use or deployment of catalysts, reactants, etc.and general process conditions are, for purposes of this invention,conventional in nature, as will readily be apparent to those skilled inthe art and in view of the above description of the invention.

It will be obvious to those skilled in the art that many modificationsmay be made within the scope of the processes of the present inventionwithout departing from the spirit thereof, and the invention includesall such modifications.

What is claimed is:
 1. A process for producing hydrocyanic acid in thegaseous phase, which comprises introducing hydrocyanic synthesis gasinto a heat exchange means; passing the resultant synthesis gas into areactor having a single-walled pressure shell, a plurality of annularcatalyst beds with particulate catalyst in each bed, and a plurality ofannular-shaped cross-flow heat exchange means, said annular catalystbeds and said annular heat exchange means being alternatingly disposed;passing said synthesis gas through the first of said beds in a radiallyoutward direction toward the pressure containment walls of said reactor;passing the effluent from said first bed through the first of saidcross-flow heat exchange means in said radially outward direction;passing the resulting effluent successively through the next bed and thenext cross-flow heat exchange means for each of the remaining beds andheat exchange means in said radially outward direction, each cross-flowheat exchange means providing inter-stage feed-effluent heat exchange,the shell-side fluid of said heat exchange means flowing radiallyoutwardly in a direction substantially normal to the direction in whichthe tube-side fluid flows; and recovering a gas which is enriched inhydrocyanic acid.